Production of biodiesel and glycerin from high free fatty acid feedstocks

ABSTRACT

A system and method for the conversion of free fatty acids to glycerides and the subsequent conversion of glycerides to glycerin and biodiesel includes the transesterification of a glyceride stream with an alcohol. The fatty acid alkyl esters are separated from the glycerin to produce a first liquid phase containing a fatty acid alkyl ester rich (concentrated) stream and a second liquid phase containing a glycerin rich (concentrated) stream. The fatty acid alkyl ester rich stream is then subjected to distillation, preferably reactive distillation, wherein the stream undergoes both physical separation and chemical reaction. The fatty acid alkyl ester rich stream is then purified to produce a purified biodiesel product and a glyceride rich residue stream. Biodiesel may be further recovered from the glyceride rich residue stream, by further separation of and/or processing of glycerides/free fatty acids contained therein. The glycerin rich second liquid phase stream may further be purified to produce a purified glycerin product and a (second) wet alcohol stream. Neutralization of the alkaline stream, formed during the alkali-catalyzed transesterification process, may proceed by the addition of a mineral or an organic acid.

SPECIFICATION

This application is a continuation-in-part application of U.S.application Ser. No. 11/504,828, filed on Aug. 15, 2006 which is acontinuation-in-part application of U.S. application Ser. No.10/766,740, filed on Jan. 26, 2004, which claims the benefit of U.S.Patent Application Ser. Nos. 60/443,049, filed Jan. 27, 2003, and60/537,251, filed Jan. 15, 2004, each of which is hereby incorporated.

FIELD OF THE INVENTION

The present invention relates to improved processes and systems forbiodiesel production.

BACKGROUND OF THE INVENTION

There is continued and growing interest in the use of renewableresources as replacements for petroleum-derived chemicals. Fatty acidalkyl esters (FAAEs) produced from fats and oils have been investigatedas replacements for such petroleum-derived materials, particularlydiesel fuel.

It has long been known that triglycerides from fats and oils can be usedas fuels for diesel engines. However, such use typically results inengine failure. Remedies for such engine failure wherein conversion offatty acids, found in lipids, into simple esters, such as methyl andethyl esters, has been proposed. See, for instance, the processdescribed in U.S. Pat. No. 6,398,707. An increasing body of evidenceindicates that these esters perform well in essentially unmodifieddiesel engines and that such esters may effectively reduce the output ofparticulate and hydrocarbon pollutants relative to petroleum-dieselfuel. The term “biodiesel” is applied to these esters.

Processes for biodiesel production have been known for many years. Forinstance U.S. Pat. No. 4,164,506 discloses a biodiesel synthesis whereinfatty acids are subjected to acid catalysis. The conversion oftriglycerides with base catalysis is described in U.S. Pat. Nos.2,383,601 and 2,494,366. Conversion of both free fatty acids andtriglycerides with enzyme catalysis is disclosed in U.S. Pat. Nos.4,956,286, 5,697,986 and 5,713,965. None of these processes, however,completely addresses the production of biodiesel from low value highfree fatty acid feedstocks.

An economic analysis of any process for the production of biodieselindicates that feedstock cost is the largest portion of production costfor biodiesel. Whereas a 15 weight percent free fatty acid (FFA)feedstock is the highest content that any contemporary commercialprocess has proposed to handle, producers (in order to conserve costs)would prefer to use feedstocks having up to 100 weight percent FFAcontent.

Further, most of the processes of the prior art are unattractive becausethey rely upon acid catalyzed esterification of fatty acids. Acidcatalysis is not suitable for processing such feedstocks containing FFAconcentrations for two principal reasons. First, an excessive amount ofacid catalyst is required in order to fully convert feedstocks havinghigh FFA content. Since the acid catalyst must be neutralized beforeprocessing the glycerides, the increased catalyst loading results in anexcessive amount of generated salt. Further, such processes generate alarge volume of waste water as disclosed in U.S. Pat. Nos. 4,303,590,5,399,731 and 6,399,800.

While enzymatic catalysis has been reported in the literature foresterification of free fatty acids, it is disadvantageous because ofreaction product inhibition from the presence of water which resultswhen the free fatty acids in the feedstock are esterified with enzymes.Another problem evidenced from enzymatic processing is the high cost ofenzymatic catalysts. Further, enzymatic catalysts have a limited life.

To avoid two-phase operation in packed bed and other reaction settings,some conventional processes for biodiesel production use volatile, toxicco-solvents. Such a process is disclosed in U.S. Pat. No. 6,642,399 B2.The use of volatile, toxic co-solvents is environmentally unacceptable.

Further, some prior art processes for producing biodiesel employ waterto wash residual glycerin and salts from the FAAEs. This, however,generates a large volume of wastewater and increases the risk of formingFAAE emulsions, as disclosed in U.S. Pat. No. 5,399,731.

To gain market share in the fuels industry, biodiesel must becompetitively priced with conventional hydrocarbon diesel. To becompetitively priced, production of biodiesel must be economicallyprofitable. Increased profitability requires that the biodiesel industrytake advantage of lower cost feedstocks. In addition, overall yields ofbiodiesel from fats and oils must be high. Increased yield is a veryimportant criterion as feedstock costs approach two thirds of the totalcost of production of biodiesel.

Improvements in processes for biodiesel production therefore need to bedeveloped which result in an increased yield of biodiesel fromfeedstocks and which minimize undesirable by-products. Alternativeprocesses further need to be developed which do not require highpressures or acid catalysis. Such processes should not employ toxicco-solvents or water for the extraction of impurities. Such processesalso need to produce high yield of biodiesel as well as employinexpensive feedstocks. Further, such feedstocks need to have a high FFAcontent in order to be competitive with petrodiesel.

SUMMARY OF THE INVENTION

A process is disclosed which combines several unit operations into aneconomical and unique process for the conversion of free fatty acids toglycerides and the subsequent conversion of glycerides to glycerin andFAAEs. The fatty acid alkyl esters of the invention produced inaccordance with the invention are typically fatty acid methyl estersthough other fatty acid alkyl esters may be produced.

The invention relates to a process for converting low-value, high freefatty acid (FFA) feedstocks to biodiesel and high quality glycerin at amarket price comparable to that of petroleum derived diesel fuels. Theprocess of the invention therefore substantially departs fromconventional concepts and designs of the background art. In so doing,the inventive process provides a process and apparatus primarilydeveloped for the purpose of producing fatty acid alkyl esters and highquality glycerin from any low-value high free fatty acid feedstock.

In a preferred aspect of the invention, streams enriched in fatty acidalkyl esters are subjected to successive treatment stages ofdistillation and/or non-evaporative separation in order to maximize theyield of recovery of purified biodiesel.

Another aspect of the invention relates to separation and purificationof major by-products of biodiesel production to render glycerin at apurity level greater than 95 or 99.7 percent, with non-detectable levelsof alcohol and less than 0.5 percent weight/weight (w/w) salts.

A further aspect of the invention relates to treatment of a by-productstream (from which biodiesel has been separated) in order to maximizethe yield of recovery of purified biodiesel.

The invention further relates to minimization of waste streams duringnormal operations, the use of lower operating conditions (such aspressures) than other commercial biodiesel processes, the non-use oftoxic co-solvents and the production of a high quality glycerinbyproduct.

In a preferred embodiment, the process is a continuous process.

The major steps of the process include the transesterification of aglyceride stream with an alcohol, preferably in the presence of basecatalyst, to convert the glycerides to fatty acid alkyl esters andglycerin.

The fatty acid alkyl esters are then separated from the glycerin toproduce a first liquid phase containing a fatty acid alkyl ester richstream and a second liquid phase containing a glycerin rich stream.

The fatty acid alkyl ester rich stream is then subjected to a firstdistillation or to a non-evaporative separation process. Preferably, thefatty acid alkyl ester rich stream is subjected to reactivedistillation, wherein the stream undergoes both separation and chemicalreaction. By means of reactive distillation, the stream is separatedinto (i.) a bottoms fraction or biodiesel stream comprising a pluralityof the fatty acid alkyl esters; and (ii.) an overhead fraction(principally composed of alcohol, a first wet alcohol stream), whilesimultaneously chemically reacting two or more stream componentstogether in such a way as to remove unwanted impurities in one or moreoutput stream(s). Such reactive distillation for example increases theyield amount of glycerides exiting the distillation column whileincreasing the purity of the biodiesel exiting the distillation column.The biodiesel exiting the distillation column may be separated into apurified biodiesel stream and a by-product stream.

The biodiesel stream exiting the first distillation column may furtherbe subjected to a second distillation or to a non-evaporative separationin order to render a purified second biodiesel stream along with asecond by-product fuel stream. The preferred second distillation occursin a wiped film evaporator or a falling film evaporator, or other suchevaporative device. Non-evaporative separation typically is a physicalseparation technique, such as freeze crystallization, steam stripping orliquid-liquid separation. A free fatty acid stream and/or glycerideenriched stream may further be separated from the second by-product fuelstream and then re-introduced into the process for production of fattyacid alkyl esters.

The glycerin rich stream of the second liquid phase may further bepurified to produce a purified glycerin product and a (second) wetalcohol stream. A portion of the purified glycerin product may then berecycled into a glycerolysis reactor (in a glycerolysis processdescribed in more detail below) for reaction with the free fatty acids.

The wet alcohol streams may further be purified, preferablycontinuously, to produce a purified alcohol product. Further, at least aportion of the purified alcohol product may be recycled into thetransesterification reactor for reaction with the glycerides.

Neutralization of the alkaline stream, formed during thealkali-catalyzed transesterification process, may proceed by theaddition of a mineral acid or more preferably an organic acid to thestream. Neutralization may occur by addition of the acid to thetransesterification effluent stream directly or to the fatty acid alkylester rich stream and/or glycerin rich stream after such streams havebeen separated from the transesterification effluent stream.

BRIEF DESCRIPTION OF THE DRAWINGS

The features of the invention will be better understood by reference tothe accompanying drawings which illustrate presently preferredembodiments of the invention. In the drawings:

FIG. 1 is a schematic flow diagram of the process of the invention.

FIG. 2 is a schematic block diagram of the biodiesel production systemin accordance with the invention;

FIG. 3 is a schematic block diagram showing the basic steps of theproduction of biodiesel in accordance with the process of the invention;

FIG. 4 is a schematic flow diagram of the process of the inventionwherein a mineral acid is used in the neutralization of the alkalicatalyst used during transesterification; and

FIG. 5 is a schematic flow diagram of the process of the inventionwherein an organic acid is used in the neutralization of the alkalicatalyst used during transesterification;

FIG. 6 is a schematic block diagram which demonstrates reactivedistillation of a fatty acid alkyl ester rich stream upon separationfrom the transesterification effluent stream, as set forth in ExampleNo. 6.

FIG. 7 is a schematic block diagram which illustrates the recycling of astream from a by-product stream for further recovery of fatty acid alkylesters.

FIG. 8 is a schematic block diagram of the process of the inventionillustrating the use of a non-evaporative separator to generate streamsenriched in fatty acid alkyl esters, glycerides and free fatty acidsfrom which refined biodiesel may be recovered.

FIG. 9 is a schematic diagram illustrating biodiesel refining wherein abiodiesel stream is treated in an evaporative device, such as a wipedfilm evaporator or falling film evaporator, for further recovery offatty acid alkyl esters.

FIG. 10 is a schematic diagram which demonstrates an embodiment of theinvention wherein by-product (fuel) separated from a biodiesel stream isfurther recycled to an evaporative device, such as a wiped filmevaporator or falling film evaporator, for further recovery of fattyacid alkyl esters.

FIG. 11 shows another embodiment of the invention wherein the by-product(fuel) stream, separated from purified biodiesel, is further separatedfor additional recovery of fatty acid alkyl esters.

FIG. 12 illustrates an embodiment of the invention wherein a biodieselstream may be directed to a non-evaporative separator, separated into afatty acid enriched stream and then re-directed to a second evaporativedevice for purification.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

In the process of the invention, biodiesel is prepared by reactingglycerides with an alcohol in a transesterification reactor to producefatty acid alkyl esters. This reaction typically occurs in the presenceof an alkali catalyst. The alcohol is typically a C₁-C₅ alcohol,preferably methanol.

The resulting transesterification effluent stream may then be separatedinto a fatty acid alkyl ester rich stream and a glycerin rich stream.Each of these streams may then be purified or subject to furtherseparation processes in order to maximize the efficiency in recovery ofbiodiesel, glycerin and alcohol. By-product (fuel) streams, separatedfrom purified biodiesel, may further be subjected to further processingin order to maximize the efficiency of biodiesel recovery.

The alkaline transesterification effluent stream formed during thealkali-catalyzed transesterification process may be directly treatedwith a neutralizing agent, such as a mineral acid or an organic acid.Alternatively, the neutralizing agent may be added to the fatty acidalkyl ester rich stream and/or the glycerin rich stream after thestreams have been separated from the transesterification effluentstream. Fatty acid alkyl esters are recovered from this pH adjustedstream.

Subsequent to neutralization, the neutralized stream may further bepurified, such as by distillation or fractionation.

The process of the invention may further consist of an esterificationstep wherein a free fatty acid feedstock is first converted toglycerides. The resulting glycerides are then introduced into thetransesterification reactor.

The use of the acid as neutralizing agent converts soaps, formed in thetransesterification reactor, to free fatty acids. The soap forms fromthe action of caustic with fatty acids in the transesterificationreactor. The presence of the soap makes it very difficult to effectuatephase separation between the fatty acid alkyl esters and the solution ofglycerin, water, alcohol and salt. As a result, the soap emulsifies andretains much of the fatty acid alkyl esters in the glycerin rich phase.Purification of the glycerin rich phase is therefore complicated by thepresence of the soap and the yield of alkyl esters is decreased.

An overview of the process of the invention may be presented in FIG. 3wherein a feedstock 1 containing free fatty acids is introduced into aglycerolysis reactor 2 with glycerin wherein the free fatty acids areconverted to glycerides. The glycerides are then introduced intotransesterification reactor 4 with alcohol wherein the glycerides aretransesterified to form fatty acid alkyl esters and glycerin.Alcohol/alkali stream 3 may be introduced into transesterificationreactor 4 as a combined mixture of alkali catalyst and alcohol, oralternatively the alkali catalyst and alcohol may be introduced into thetransesterification reactor as separate streams into transesterificationreactor 4. The transesterification effluent stream 4 a or a portionthereof is then neutralized during neutralization/phase separation step5, either before or after the effluent stream 5 a is separated into afatty acid alkyl ester rich stream and a glycerin rich stream.Ultimately, alcohol, glycerin and biodiesel may be refined in alcoholrefining step 6, glycerin refining step 7 and biodiesel refining step 8,respectively. The alcohol typically exits the system as a small portionof waste stream 9 a or is recycled via flow 11 back to thetransesterification reactor. Refined glycerin is isolated in technicalgrade glycerin stream 13 and/or may be recycled back via flow 15 toglycerolysis reactor 2. Waste stream 9 b may contain some unrefinedglycerin. The alkyl esters may further be refined in biodiesel refiningstep 8 to produce purified biodiesel stream 18 and waste stream 19 whichmay be useful, for example, as a burner fuel.

As illustrated in FIG. 7, at least a portion of the waste stream 19 maybe reintroduced into prior processes, for example as stream 351, intobiodiesel refining stage 8 to further recover fatty acid methyl esters,or into the transesterification reactor 4 to transesterify glyceridesinto fatty acid methyl esters, or into esterification reactor 2 toesterify fatty acids.

Alternatively, as illustrated in FIG. 8, at least a portion 351 of wastestream 358 may be separated into (i.) fatty acid alkyl ester enrichedstream 371 and (ii.) glyceride enriched stream 376 and/or free fattyacid enriched stream 374 in separator 370. Fatty acid alkyl esterenriched stream 371 may then be re-introduced into biodiesel refiningstage 8. The glyceride 376 and/or free fatty acid 374 enriched streamsmay then be re-introduced into the transesterification reactor 4 and/oresterification reactor 2.

The process of the invention may be a continuous process. For example, acontinuous process, wherein one or more of the following steps arecarried out in a continuous fashion, is apparent from the descriptionprovided herein:

(1) the optional conditioning of a fatty acid containing feedstock byheating, mixing and filtering;

(2) continuously reacting the free fatty acids in the feedstock withglycerin in a glycerolysis or esterification reactor to produceglycerides;

(3) reacting the glycerides in a transesterification reactor withalcohol to render fatty acid alkyl esters and glycerin. This reactionpreferably occurs in the presence of an alkali catalyst;

(4) separating (e.g., by gravitational separation of two relativelyimmiscible phases), fatty acid alkyl esters and glycerin from thetransesterification effluent stream to yield a fatty acid alkyl esterrich stream and a glycerin rich stream;

(5) purifying the fatty acid alkyl ester rich stream by distillationand/or fractionation. In a preferred embodiment, the fatty acid alkylester rich stream is purified by reactive distillation wherein areaction in the distillation or fractionation column assists in thereduction of unwanted impurities such as glycerin. The purified fattyacid alkyl ester is acceptable for use as biodiesel;

(6) purifying the glycerin rich stream, preferably by use of an organicacid, such as a weak organic acid like acetic acid, formic acid orpropionic acid, and recovering alcohol from the stream. The purifiedglycerin may then be introduced into the glycerolysis reactor;

(7) purifying the wet alcohol streams resulting from steps (5) and (6)above and removing water from the streams; and

(8) recycling at least a portion of the purified alcohol to thetransesterification reactor for reaction with the glyceride.

The process may further consist of subjecting the biodiesel stream ofstep (5) to further separation by a second distillation ornon-evaporative separation in order to render a more purified biodieselstream (or second purified biodiesel stream) and a second by-productfuel stream.

As another option, the biodiesel stream of step (5) may further beseparated in a non-evaporative separator into (i) a fatty acid alkylester enriched stream and (ii) a glyceride and/or free fatty acidenriched stream. Preferred non-evaporative separators for use hereinclude freeze crystallization processes and liquid-liquid separationprocesses.

The fatty acid alkyl ester enriched stream, resulting from thisseparation, may then be combined with the biodiesel stream of step (5)and then subjected to the second distillation or non-evaporativeseparation. The glyceride and free fatty acid enriched stream may thenbe re-introduced to the transesterification or esterification reactors.

The feedstock, from which the biodiesel may be produced, typicallycontains a plurality of free fatty acids. The feedstock typicallycontains between from about 3 to about 100 weight percent of free fattyacids and, optionally, a fat and/or oil.

Typically, the feedstock is a lipid feedstock. The free fatty acidfeedstock for use in the invention may be a low-grade lipid materialderived from animal fats and vegetable oils, including recycled fats andoils. For instance, the feedstock for the production of biodiesel fuelmay be a grease feedstock, such as a waste grease or a yellow grease.Such low-grade lipid materials are very complex and typically aredifficult to economically process using current state of the artprocesses because of their high free fatty acid levels (ranging from afew percent to 50 percent, and higher). In addition, such materialscontain unprocessable material and contaminants that must be removedprior to processing or during refinement of the products.

The feedstock may be first introduced into a conditioning vessel orreactor that is operative to heat, mix and/or filter the feedstock toproduce a conditioned feedstock. The feedstock may then be filtered,such as by using a traveling screen.

Subsequent to filtration, the concentration of free fatty acids in theconditioned lipid feedstock may be measured. Optionally, theconcentration of free fatty acids in the conditioned feedstock may becontinuously measured throughout the process. Measurements may be madewith an in-line free fatty acid measurement device, such as a titrationdevice or near-infrared spectrophotometer, that is operative to quantifythe concentration of the free fatty acid in the conditioned feedstock.

During conditioning, the feedstock may be heated to a temperature in therange of about 35° C. to about 65° C., preferably between from about 55°C. to about 65° C., while mixed. A uniform mixture of glycerides, freefatty acids and unsaponifiable materials are typically present in theconditioned feedstock.

During glycerolysis, glycerin is used as a reactant to convert the freefatty acids in the feedstock to glycerides (mono-, di-, andtriglyceride). Reaction of the free fatty acids in the feedstocktypically occurs in the absence of a catalyst. In the glycerolysisreactor, the free fatty acid in the feedstock is mixed and continuouslyreacted with glycerin at an appropriate temperature and pressure torender a glycerolysis reactor effluent stream that contains generallyless than about 0.5 percent by weight of free fatty acids and aplurality of glycerides. Glycerolysis preferably occurs in the absenceof both catalyst and co-solvent.

The glycerin, typically a purified glycerin product, is normally addedto the glycerolysis reactor at a rate that is greater than thestoichiometric amount of glycerin required for the glycerolysisreaction. The amount of glycerin introduced to the glycerolysis reactoris generally in a stoichiometric proportion of about 35 percent to about400 percent glycerin to free fatty acid in order to render theglyceride. In a preferred embodiment, the amount of glycerin added tothe glycerolysis reactor is at a rate in the range of about 300 percentof the stoichiometric amount of free fatty acids in the feedstock.

Preferably, glycerolysis is conducted at a temperature in the range ofabout 150° C. to about 250° C., typically from about 180° C. to about250° C., more typically from about 180° C. to 230° C. The reactiontypically proceeds under agitation. The reaction is further typicallyconducted at a pressure of about 0.1 pounds per square inch absolute toabout 15 pounds per square inch absolute, more typically about 2 poundsper square inch absolute.

Reaction of the free fatty acids and glycerin typically occurs in thepresence of a catalyst such as ZnCl₂, but in a preferred embodiment isperformed in the absence of a catalyst. The glycerolysis reactoreffluent stream may contain less than 0.5 percent by weight of freefatty acids and a plurality of glycerides.

The glycerolysis is typically a continuous reaction. The continuousreaction of the free fatty acid in the feedstock with glycerin toproduce the glyceride in the glycerolysis reactor may be conducted inresponse to a signal from the in-line fatty acid measurement device orspectrophotometer.

During glycerolysis, water is removed; the produced glycerides beingessentially water-free. Water is typically continuously removed from theglycerolysis reactor as a vapor through a fractionation column or a ventin the reactor headspace. Preferably, the vapor vented from theglycerolysis reactor is fractionated to yield three streams, the firstfraction having a high concentration of unsaponifiables evaporated fromthe feedstock that are condensed as a liquid stream, the second fractionbeing a liquid fraction having a high concentration of glycerin, and avapor fraction and a third liquid fraction having a high concentrationof water. The liquid fraction containing the glycerin may then bereturned to the glycerolysis reactor.

The glycerolysis reactor may consist of two or more continuous stirredtank reactors operated in series. The residence time of such reactors istypically from about 30 to not more than about 500 minutes, andpreferably not more than 200 minutes.

A plurality of glycerides contained in the glycerolysis effluent streamis reacted with an alcohol in the transesterification reactor, such as acontinuous stirred tank reactor. In this reaction, the glycerides in theglycerolysis reactor effluent stream are transesterified into fatty acidalkyl esters and glycerin. Transesterification proceeds at anappropriate temperature and pressure to produce the desiredtransesterification reactor effluent stream.

Transesterification, which preferably is a continuous process, occurs inthe presence of a base catalyst. Suitable base catalysts include suchalkali catalysts as potassium hydroxide and sodium hydroxide. The alkalicatalyst may be added to the transesterification reactor at a ratesufficient to catalyze the reaction. Typically, the amount of alcoholadded to the transesterification reactor is from about 1 mole to 5 molesalcohol to each mole of fatty acid portion of the glycerides present inthe transesterification reactor inlet stream. More typically, the ratiois about 2 moles alcohol for each mole of fatty acid portion present inthe glycerides introduced into the transesterification reactor. Thecatalyst, typically potassium hydroxide, is added at a ratio of about0.5% to 3% by weight catalyst to weight glycerides, more typically about1%.

Alternatively, an alkoxide, such as potassium methylate, may be added tothe transesterification reactor to facilitate the base catalysis. Assuch, the rapid conversion of glycerides to alkyl esters may occur inthe presence of caustic alkoxide, such as caustic methoxide catalysts.

The transesterification reaction typically occurs at a temperature inthe range of about 25° C. to about 65° C., preferably from about 50° C.to about 60° C., and at a pressure of about 14.5 psia to about 3,625psia.

The alcohol is normally added to the transesterification reactor at arate that is greater than the stoichiometric amount of alcohol requiredfor the alkali catalyzed transesterification reaction. For instance, thealcohol may be added to the transesterification reactor at a rate equalto about 200 percent of the stoichiometric amount of alcohol requiredfor the catalyzed reaction.

Preferably, multiple alcohol or catalyst additions are made to thetransesterification reactor.

The transesterification reactor typically contains at least twocontinuous stirred tank reactors that are operated in series. Each ofthe tank reactors typically has a residence time of about 5 minutes toabout 90 minutes, typically about 60 minutes.

The resulting transesterification reactor effluent stream contains afatty acid alkyl ester and glycerin. Preferably, at least a portion ofthe glycerin is removed from the transesterification reactor before theplurality of glycerides is reacted with the alcohol.

A plurality of the resulting fatty acid alkyl esters may then beseparated from the glycerin in the transesterification effluent stream.Separation into two distinct immiscible phases, i.e., a first liquidphase in which the plurality of fatty acid alkyl esters may beconcentrated and a second liquid phase in which glycerin may beconcentrated, is typically dependent upon the differences in densitiesin the two phases and employs gravitational force and/or centrifugalforce.

Typically, the two phases are separated at a temperature of about 25° C.to about 65° C. to produce the fatty acid alkyl ester rich stream andglycerin rich stream. This separation process may be a continuousoperation and may be performed in a clarifier or by means of membranefiltration.

In a preferred embodiment, the fatty acid alkyl ester rich stream issubjected to reactive distillation in biodiesel refining 8 to separatethe fatty acid alkyl ester rich stream into a bottoms fraction, anoverhead fraction (principally comprising excess alcohol) and a fattyacid alkyl ester product stream. Such separation utilizes thedifferences in the vapor pressures of the components of the fatty acidalkyl ester rich stream and the reactive loss of glycerin. Theconditions in the distillation or fractionation column includingtemperature and pressure conditions, simultaneously with and in the samevessel wherein the said separation occurs, promote a chemical reactionto occur. Reactive distillation in the embodiment depicted in FIG. 6decreases the concentration of glycerin and increases the amount ofglycerides exiting the column. Thus, reactive distillation increases theefficiency of the production process.

The end result of reactive distillation is that the amount of glycerinseen in the transesterification effluent stream, or the first liquidphase, is greater than the total amount of glycerin which exits thedistillation or fractionation column. This is attributable to thereaction of the glycerin with free fatty acids and or fatty acid alkylesters in the reactive distillation column to form glycerides.

Preferably, the overhead fraction produced by the fatty acid alkyl esterdistillation column is a (first) alcohol stream which comprisesessentially the alcohol. Preferably the bottoms fraction comprisesimpurities having a high boiling point, unsaponifiable materials,monoglycerides, diglycerides, triglycerides and fatty acids.

Preferably, the fatty acid alkyl ester distillation column orfractionation column is operated at a pressure below about 15 pounds persquare inch absolute. More preferably, the fatty acid alkyl esterdistillation column or fractionation column is operated at a pressure inthe range of about 0.1 pounds per square inch absolute to about 3 poundsper square inch absolute. Preferably, the fatty acid alkyl esterdistillation column or fractionation column is operated at a temperaturein the range of about 180° C. to about 290° C., more preferably betweenfrom about 230° C. to about 270° C. Preferably, the fatty acid alkylester distillation column or fractionation column contains a packingmaterial.

The glycerin rich second liquid phase stream may further be purified andalcohol recovered from it. The recovered alcohol is operative to producea purified glycerin product and a (second) wet alcohol stream. In apreferred embodiment, this step employs one or more of glycerinfractionation (wherein the fractions within the glycerin rich stream areseparated by distillation), phase separation (wherein the impuritiesthat co-fractionate with glycerin are removed by immiscibility anddifferences in density) and glycerin polishing (wherein other impuritiesare removed from glycerin).

The glycerin rich stream may further be subjected to phase separationwherein a fatty acid alkyl ester rich liquid phase and a glycerin richliquid phase are separated and the two liquid phases may then be subjectto purification as described in the paragraphs above.

The glycerin rich stream may further be purified in a glycerindistillation or fractionation column to produce a bottoms material, aside stream and an overhead stream. Preferably, the bottoms materialcontains essentially waste materials; the side stream containsessentially glycerin and trace impurities; and the overhead streamcontains essentially alcohol and water that is collected for furtherpurification and recycled.

Preferably, the glycerin distillation column is operated at an elevatedtemperature between about 180° C. and about 280° C., more preferablybetween from about 180° C. to about 230° C. The distillation column istypically operated at a reduced pressure, of below about 2 pounds persquare inch absolute, typically the pressure is in the range of about0.1 pounds per square inch absolute to about 2 pounds per square inchabsolute.

The glycerin rich stream may further be subjected to a decolorizationcolumn wherein colored impurities and odors are removed from theglycerin, i.e., “glycerin polishing”. The decolorization columntypically comprises a packed bed of activated carbon operated at atemperature in the range of about 35° C. to about 200° C., preferablybetween from about 40° C. to about 100° C. The contact time is generallyless than four hours. Activated carbon fines carried through the packedbed are removed by filtration.

Water may further be removed from the wet alcohol streams to renderpurified alcohol by subjecting the wet alcohol stream to an alcoholdistillation or fractionation column at a temperature in the range ofabout 60° C. to about 110° C. and at a pressure in the range of about 14pounds per square inch absolute to about 20 pounds per square inchabsolute. Preferably, this purification comprises adsorption ontomolecular sieves that can then be dried and reused or distillationresulting in a bottoms product consisting mainly of water.

At least a portion of the purified glycerin product may then be returnedto the glycerolysis reactor for reaction with free fatty acids in thefeedstock; at least a portion of the purified alcohol being recycledinto the transesterification reactor for reaction with glycerides.

It is typically desired to neutralize the fatty acid alkyl ester andglycerin produced in the transesterification reactor. Neutralization isoften required in light of the caustic conditions which characterizetransesterification. Such neutralization may occur by addition of anacid to the transesterification effluent stream or to either the fattyacid alkyl ester rich stream or glycerin rich stream after such streamsare separated from the transesterification effluent stream. Suitableacid treatments include mineral or more preferably organic acidtreatments.

Suitable mineral acids include sulfuric acid and phosphoric acid.Reaction of the alkali catalyst with a mineral acid renders an insolublesalt that is removed from the glycerin rich stream in a solidsseparation operation.

FIG. 4 is illustrative of the process wherein a mineral acid, such asphosphoric acid, is employed. In particular, FIG. 4 illustratesintroduction of feedstock 1 containing free fatty acids intoglycerolysis reactor 2 wherein the free fatty acids are converted toglycerides by esterification. The glycerides are then introduced intotransesterification reactor 4 with alcohol 3 and alkali catalyst 318 at317 (illustrated in FIG. 7) wherein the glycerides are transesterifiedto form fatty acid alkyl esters and glycerin.

The transesterification effluent stream 4 a is first separated in 1^(st)phase separation 320, typically by gravitational separation techniques,into a fatty acid alkyl ester rich stream and a glycerin rich stream.Each of these streams may then be purified in 2nd phase separation 322in accordance with the processes described herein.

The neutralization acid, phosphoric acid, 324 is added either prior to1^(st) phase separation 320 or subsequent to 1^(st) phase separation 320of the transesterification effluent stream after the fatty acid alkylester rich stream and glycerin rich stream have been separated. Suchalternative or combination ports of introduction of the acid into theprocess are represented by the dotted lines in FIG. 4.

Unfortunately, use of phosphoric acid renders an insoluble precipitate.The formation of the insoluble precipitate mandates the use of a filterin filtration step 326 and/or a filter in filtration step 328. Suitablefilters include rotary vacuum drum filters, plate and frame presses aswell as belt presses.

In addition to the use of a filtration unit, use of a mineral acidfurther requires the rinsing of the insoluble by-product salts in orderto wash residual organic materials from them. Suitable solvents includeC₁-C₅ alcohols, such as methanol. Illustrated in FIG. 4 is theintroduction of alcohol solvent 329 for use as alcohol rinse 330 whichremoves organic residue from the filter cake. Vacuum dry 332 is thenused to remove alcohol from the filter cake and to dry the purified saltwhich then exits the process as waste stream 334. The solvent may thenbe recovered as stream 364 for reuse in the process.

Preferably, the process comprises drying the insoluble salt in a dryerunder conditions wherein the temperature of the dryer exceeds theboiling point of the solvent at the operating pressure of the dryer. Thedryer may optionally be operated under a vacuum to improve the drying.The dryer may further include a condenser to recover the solvent forreuse.

FIG. 4 further illustrates the refining of alcohol, glycerin andbiodiesel in alcohol refinery vessel 6, glycerin refinery vessel 7 andbiodiesel refinery vessel 8, respectively. The alcohol typically exitsthe system as byproduct stream 9 a or is recycled via 11 back totransesterification reactor 4. Refined glycerin is isolated as purifiedglycerin 13. A portion of the glycerin stream may be recycled back asstream 15 to glycerolysis reactor 2. The alkyl esters may further bepurified to produce purified biodiesel 18 or may exit the system asbyproduct 19 in the form of, for example, burner fuel.

It is more preferable to employ an organic acid versus a mineral acid,however. While there are inorganic acids that don't create precipitatingsalts upon neutralization with the transesterified stream, all sufferfrom serious disadvantages. For instance, hydrochloric and perchloricacid produce chlorides in the process streams which, in turn, causeundesirable corrosion of steel and stainless steel, especially atelevated temperatures. Sulfuric acid, sulfurous acid and hydrogensulfide suffer serious disadvantages due to the presence of sulfur whichincreases the tendency of sulfur to exit with the final biodieselproduct. This, in turn, causes potential failure of sulfur level limitsand the formation of unwanted sulfur oxide in emissions frombiodiesel-burning engines. Arsenic acid, chromic acid, hydrocyanic acidand hydrofluoric acid are undesirably hazardous to use and/or requireunwanted additional treatment methods for the disposal of undesirablebyproducts. Lastly, iodic acid does not produce undesirableprecipitates, but it is economically not viable.

When an organic acid is used, no insoluble salt is formed and thus it isunnecessary to subject the stream to any solids separation operation.Suitable organic acid include weak organic acids, such as formic acid,acetic acid and propionic acid. In such instances, the pH of theglycerin rich stream resulting from transesterification may first beadjusted below 8.0, preferably between from about 6.5 to about 7.0.

FIG. 5 contrasts the inventive process wherein an organic acid 325 isused in the neutralization of the alkali catalyst versus a mineral acid.In one embodiment, organic acid is added to the transesterificationeffluent prior to separation of the fatty acid alkyl ester rich streamform the glycerin rich stream, at a weight ratio of from about 0.1% toabout 5%, more typically about 0.9%. In another embodiment, organic acidis added to the glycerin rich stream at a weight ratio of from about 1%to about 7%, more typically about 4%. The use of an organic acid rendersthe steps of filtration, rinsing of the filter cake and vacuum dryingunnecessary and thus offers advantages over the use of the mineral acid.

As illustrated in FIG. 7, a portion of by-product (fuel) stream 351 isshown as being directed back into biodiesel refining stage via 351A,into transesterification reactor via 351C, or into esterificationreactor via 351D. The composition of stream 351 is not changed prior tobeing separated into streams 351A, 351C and 351D.

In contrast, in FIG. 8, a portion of by-product (fuel) stream 351 isseparated in separator 370 into fatty acid alkyl ester enriched stream371 and/or a second stream 374 enriched in free fatty acids and/or athird stream 376 enriched in glycerides. The portion of the secondstream having the lower free fatty acid content is then introduced intotransesterification reactor 4 and the portion of stream having thehigher free fatty acid content is introduced into esterification reactor2.

FIG. 9 illustrates an embodiment for biodiesel refining step 8 whereinan increased yield of biodiesel may result by the use of a seconddistillation reactor or non-evaporative separator. In a preferredembodiment, this second distillation reactor is one or more evaporativedevices, such as wiped film evaporators or falling film evaporatorsknown in the art. Typically, this second distillation reactor occurs inthe biodiesel refining unit. Further, a separator unit may also be usedto treat the by-product (fuel) stream which results from thepurification of biodiesel.

A system may be constructed in accordance with the teachings set forthherein for the production of biodiesel from a feedstock, such as a lipidfeedstock having free fatty acids. The system may include:

(1) an optional conditioning reactor which is operative to continuouslyconvert the feedstock to a conditioned feedstock. The conditioningreactor is operative to heat, mix and filter the feedstock in order toproduce a conditioned feedstock;

(2) an optional system for continuously measuring the concentration ofthe free fatty acid in the conditioned feedstock. Suitable systemsinclude an in-line free fatty acid measurement device which is operativeto quantify the concentration of the free fatty acid in the conditionedfeedstock;

(3) a glycerolysis reactor wherein the free fatty acid in the feedstockis continuously reacted with glycerin to produce a glyceride. Thisreaction may be in response to a signal from the in-line free fatty acidmeasurement device;

(4) a transesterification reactor for continuously reacting theglyceride with an alcohol and which is operative to convert theglyceride to a fatty acid alkyl ester and glycerin, preferably by analkali catalyzed reaction. This reaction may proceed in response to thesignal from the in-line free fatty acid measurement device;

(5) a separator for continuously separating the fatty acid alkyl esterfrom the glycerin and which is operative to produce a fatty acid alkylester rich stream and a glycerin rich stream. Suitable separatorsinclude a clarifier or a phase separation centrifuge which is operativeto produce a (first) liquid phase in which the fatty acid alkyl ester isconcentrated and a (second) liquid phase in which glycerin isconcentrated.

(6) a purifier for continuously purifying the fatty acid alkyl esterrich stream and recovering the alcohol from the fatty acid alkyl esterrich stream; the purifier being operative to produce a purifiedbiodiesel product and a first wet alcohol stream. Suitable purifiersinclude fractionation and distillation columns. In a preferredembodiment, the fatty acid alkyl ester rich stream is purified byreactive distillation to render biodiesel;

(7) an optional evaporator separator, such as a wiped film evaporator ora falling film evaporator, for further separation of biodiesel into afatty acid alkyl ester enriched stream and a by-product stream (fuel)stream;

(8) an optional non-evaporative separator for separation of theby-product (fuel) stream into a fatty acid alkyl ester enriched streamand a free fatty acid/glyceride enriched stream;

(9) a purifier for continuously purifying the glycerin rich stream andrecovering alcohol from the glycerin rich stream; the purifier beingoperative to produce a purified glycerin product and a second wetalcohol stream. Suitable purifiers include fractionation anddistillation columns, including reactive distillation;

(10) a purifier for continuously purifying the wet alcohol streams thatis operative to produce a purified alcohol product. Suitable purifiersinclude an alcohol fractionation column for treating the alcoholstreams; and

(11) pathways for recycling at least a portion of the purified glycerinproduct to the glycerolysis reactor and recycling at least a portion ofthe purified alcohol into the transesterification reactor forcontinuously reacting with the glyceride.

Referring to FIG. 1, a preferred embodiment of a biodiesel productionprocess 10 for the conversion of high free fatty acid feedstocks intobiodiesel is presented.

In feedstock introduction step 12, feedstock is introduced to process10. The introduced feedstock is preferably conditioned in feedstockconditioning operation 14 wherein feedstock is heated and mixed inconditioning reactor 16; the high free fatty acid feedstock being heatedand mixed to ensure a uniform mixture. The free fatty acid may bequantified, such as in an in-line free fatty acid measurement device 18,wherein the concentration of free fatty acids in the feedstock isdetermined by spectroscopy, titration or other suitable means. In afirst separation, solid (insoluble) substances are removed in filter 24.

The feedstock may include at least one free fatty acid at aconcentration in the range of about 3 percent to about 97 percent byweight; moisture, impurities and unsaponifiable matter at aconcentration up to about 5 percent by weight; and a remainder thatincludes monoglycerides, diglycerides and/or triglycerides. Thefeedstock may further include trap grease.

Preferably, the conditioning step is carried out and produces aconditioned feedstock with a temperature in the range of about 35° C. toabout 250° C. and more preferably in the range of about 45° C. to about65° C. In a preferred embodiment, the feedstock is heated to atemperature in the range of about 55° C. to about 65° C. Preferably, theresulting conditioned feedstock is substantially free of insolublesolids.

The conditioned feedstock is introduced to a glycerolysis oresterification reaction at 26 which preferably comprises glycerinaddition step 28, heating step 32, glycerolysis step 34 in which freefatty acids are converted to glycerides and glycerolysis effluentcooling step 38.

Preferably, glycerolysis reaction step 26 further comprises performingthe glycerolysis reaction at a temperature in the range of about 150° C.to about 250° C.; and removing water from the environment of theglycerolysis reaction. More preferably, glycerolysis reaction step 26further comprises using two or more continuous stirred tank reactors inseries.

In a preferred embodiment the free fatty acid and glycerin arecontinuously reacted, typically in the absence of a catalyst, in aglycerolysis reactor at a temperature of about 220° C. and at a pressureof about 2 pounds per square inch absolute, in an esterificationreaction to produce an effluent stream that contains less than 0.5percent by weight of free fatty acids and a plurality of glycerides.Preferably, the purified glycerin product is continuously added to theglycerolysis reactor at a rate in the range of about 35 percent to about400 percent of the stoichiometric amount of free fatty acids and wateris continuously removed from the glycerolysis reactor as a vapor inwater venting step 35 through a fractionation column that returnscondensed glycerin to the glycerolysis reactor.

Preferably, the reactor for glycerolysis step 34 comprises at least twocontinuous stirred tank reactors that are operated in series, thereactors having a combined residence time of not greater than about 400minutes for feedstock with a 20 percent by weight free fatty acidconcentration.

Water is preferably removed as vapor through a fractionation column or adistillation column that returns condensed glycerin to the glycerolysisreactor.

The effluent from glycerolysis reaction step 26 is introduced to alkalicatalyzed transesterification reaction at 42 which preferably comprisesalcohol metering step 44, catalyst metering step 46, alkoxide additionstep 48 and transesterification step 50 wherein the glycerides undergotransesterification in the transesterification reactor.

In transesterification step 50, glycerides are contacted with aneffective amount of alcohol and an effective amount of alkali catalystunder conditions wherein the glycerides, alcohol and alkali catalystcome into substantially intimate contact. Preferably, the alkalicatalyst is selected from the group consisting of sodium hydroxide andpotassium hydroxide.

The transesterification reaction step 42 is preferably conducted at atemperature in the range of about 20° C. to about 65° C. and at anabsolute pressure in the range of about 14.5 psia. More preferably,transesterification reaction step 42 comprises conducting thetransesterification at a temperature in the range of about 25° C. toabout 65° C. and at an absolute pressure near atmospheric. In apreferred embodiment, the alcohol and alkali catalyst are mixed atprescribed rates prior to their addition to the transesterificationreaction operation.

In a preferred embodiment, transesterification reaction step 42comprises reacting the plurality of glycerides contained in theglycerolysis effluent stream with an alcohol in the transesterificationreactor. In the transesterification reactor, the plurality of glyceridesare preferably mixed with the alcohol and alkali catalyst by an agitatorand continuously reacted with the alcohol.

Preferably, the alcohol, most preferably methanol, is added to thetransesterification reactor at a rate equal to about 200 percent of thestoichiometric amount of alcohol required for the catalyzed reaction andthe alkali catalyst is added to the transesterification reactor at arate of about 0.5 percent by weight to 2.0 percent by weight ofglycerides present in the glycerolysis effluent stream. More preferably,the alkali catalyst is dissolved in the alcohol prior to theirintroduction to the transesterification reactor.

Preferably, the transesterification reactor comprises at least twocontinuous stirred tank reactors that are operated in series, saidreactors having a combined residence time of not more than about 90minutes.

The transesterification reactor effluent stream contains a plurality offatty acid alkyl esters and glycerin. The effluent fromtransesterification reaction step 42 is preferably introduced to secondseparation at 52 in which a light phase (for instance, specific gravity0.79-0.88) is separated from a heavy phase (for instance, specificgravity 0.90-1.20). In biodiesel purification step (operation) 58(referenced as 8 in FIG. 3), excess methanol and high-boiling impuritiesare preferably separated from fatty acid alkyl esters in the light phaseand the alcohol is collected for reuse. Preferably, separating the fattyacid alkyl esters from the glycerin involves using the densitydifference between the first light liquid phase and the second heavyliquid phase to separate them.

In biodiesel purification step 56, differences in component vaporpressures are used to separate excess alcohol and high-boilingimpurities from fatty acid alkyl esters in the light phase, and thealcohol is collected for reuse.

In a preferred embodiment, second separation step 52 comprisesseparating the fatty acid alkyl esters from the glycerin in thetransesterification effluent stream in a continuous clarifier in phaseseparation step 54. Preferably, in the continuous clarifier, a firstlight liquid phase in which the plurality of fatty acid alkyl esters areconcentrated and a second heavy liquid phase in which glycerin isconcentrated are continuously separated at a temperature of about 25° C.to about 65° C. to produce a fatty acid alkyl ester rich stream and aglycerin rich stream.

Alternatively, the separation step may be a reactive distillation orfractionation column wherein the fatty acid alkyl ester and glycerin maybe separated. The transesterification effluent stream entering thereactive column contains, in addition to fatty acid alkyl esters, acertain amount of glycerin, glycerides and unreacted or non-convertiblelipid feedstock. In the reactive column, some of the glycerin reactswith unreacted fatty acids and/or fatty acid alkyl esters to formglycerides.

In preferred embodiments, the light phase is separated in fatty acidalkyl esters purification step 56. In step 56, differences in componentvapor pressures are used to separate excess alcohol and high-boilingimpurities from fatty acid alkyl esters in the first liquid phase, andthe alcohol is collected for reuse.

Preferably, purifying the fatty acid alkyl ester rich stream step 58further comprises using a distillation column to separate the fatty acidalkyl ester rich stream into a bottoms fraction, an overhead fractioncomprising primarily the alcohol, and a side stream fraction comprisinga fatty acid alkyl ester product. Preferably, the bottoms fractionproduced by the distillation column comprises impurities, unsaponifiablematerials, monoglycerides, diglycerides, triglycerides and free fattyacids. Preferably, the fatty acid alkyl ester product produced by thedistillation column meets ASTM specification D 6751. Preferably, theoverhead fraction produced by the distillation column comprisesessentially the alcohol.

In preferred embodiments, the heavy phase from second separation step 52is treated in catalyst separation step 62 comprising mineral acidaddition step 64, catalyst precipitation step 66 in which the alkalicatalyst is reacted with a mineral acid to produce a solid precipitate,catalyst precipitation reactor effluent filtration step 70 in which analcohol washing step 68 occurs before the alkali salt precipitate isremoved in salt recovery step 71, filtrate separation step 72 in whichthe precipitate-free filtrate is separated into two liquid phases, withthe fatty acids and fatty acid alkyl esters floating to the top and theglycerin and most of the alcohol sinking to the bottom, pHneutralization step 74 in which the pH of the glycerin is increased, andfree fatty acid recycling step 76.

Crude glycerin may be treated in glycerin purification step 80 whereinglycerin is purified by differences in component vapor pressures. Apreferred embodiment comprises distillation or fractionation step 84 inwhich the alcohol and high boiling impurities are separated from theglycerin. Glycerin decolorization step 86 comprises using a packed bedof activated carbon to remove color and odor from the distilledglycerin.

Preferably, in purifying the glycerin rich stream and recovering alcoholfrom it to produce the purified glycerin product and a wet alcoholstream, the alkali catalyst in the glycerin rich stream is reacted witha mineral acid, such as phosphoric acid or sulfuric acid, to produce aninsoluble salt having fertilizer value that is removed from the glycerinrich stream in a solids separation operation and thereafter filtered andrinsed with the alcohol.

The pH of the glycerin rich stream is adjusted to about neutral byadding a caustic alkali solution and then further purified in a glycerindistillation column that is operated at a temperature in the range ofabout 180° C. to about 230° C. and at a pressure below about 1 pound persquare inch absolute and in a decolorization column comprising a packedbed of activated carbon operated at a temperature in the range of about40° C. to about 200° C.

In a more preferred embodiment, the pH of the glycerin rich stream isadjusted to between about 6.5 and 8.0 by the addition of an acid. Anorganic acid, such as a weak organic acid, like acetic acid, propionicacid or formic acid, is then introduced to the glycerin rich stream.Salts present in the glycerin rich stream remain soluble. Thus,filtering and rinsing steps are unnecessary by use of the organic acid.

Preferably, the wet alcohol is treated in alcohol purification step 88in which water is removed from the wet alcohol. More preferably, thewater is removed by vapor pressure differences or adsorption. In apreferred embodiment, the alcohol is purified by distillation orfractionation in alcohol distillation or fractionation step 90. In apreferred embodiment, purifying the wet alcohol stream comprisesremoving water from it to produce a purified alcohol product.Preferably, the wet alcohol stream is purified in an alcoholdistillation column that is operated at a temperature in the range ofabout 60° C. to about 110° C. and at a pressure in the range of about 14pounds per square inch absolute to about 20 pounds per square inchabsolute.

In glycerin recycling step 92, glycerin is preferably recycled to step28 and in alcohol recycling step 94, alcohol is preferably recycled tostep 44. Preferably, glycerin recycling step 92 involves recycling atleast a portion of the purified glycerin product into the glycerolysisreactor for reaction with the plurality of free fatty acids in thefeedstock. Preferably, the alcohol recycling step involves recycling atleast a portion of the purified alcohol product into thetransesterification reactor for reaction with the plurality ofglycerides. The additional alcohol required for the transesterificationreaction is supplied to the alkoxide tank. Biodiesel is delivered to itsmarket in biodiesel delivery step 96 and glycerin is delivered to itsmarket in glycerin delivery step 98.

Referring to FIG. 2, a preferred embodiment of system 110 for theconversion of high free fatty acid feedstocks into biodiesel ispresented. Biodiesel production system 110 preferably comprises thesubsystems and reactors described below wherein the alcohol employed ismethanol.

In feedstock introduction subsystem 112, the feedstock is introduced tosystem 110. In a preferred embodiment, the feed material is composed ofbetween 0 and 100 percent free fatty acid content, with the remaindercomprising mono-, di- and triglycerides, moisture, impurities andunsaponifiables (MIU).

The introduced feedstock may optionally be conditioned in feedstockconditioning subsystem 14 comprising feedstock heating and mixing vessel16 in which the high free fatty acid feedstock is heated and mixed toensure a uniform, homogeneous mixture with uniform viscosity. Theconcentration of free fatty acids in the feedstock may be measured byin-line measurement device 18. The concentration is measuredcontinuously to allow continuous control of downstream process steps.

Preferably, the feed material is heated in feedstock heating and mixingvessel 16 to ensure that all of the available lipids are liquid and thatsolids are suspended. Temperatures in the range of at least 35° C. butnot more than 200° C. are adequate to melt the lipids, decrease theirviscosity and allow thorough mixing of the feedstock. A jacketed stirredtank may be used to provide agitation and maintain the feedstock atincreased temperature.

The conditioned feedstock may then be introduced to glycerolysisreaction subsystem 26 which comprises glycerin addition apparatus 28,input heater 32, first glycerolysis reactor 134 and second glycerolysisreactor 136 and glycerolysis effluent cooler 38. The filtered product ofstep 24 is combined with glycerin and subjected to conditions thatpromote the glycerolysis reaction in glycerolysis reaction subsystem126. In a preferred embodiment, these conditions include a reactiontemperature between from about 150° C. to about 250° C. and a pressurebetween about 0.1 pounds per square inch, absolute (psia) and about 30psia. A more preferred condition is a temperature of about 220° C. and apressure of about 2 psia.

Glycerin is added to the filtered grease feedstock in excess of the freefatty acid molar quantity of the grease feedstock. This excess is in therange of 10 percent to 300 percent excess glycerin (from 110 percent to400 percent of the stoichiometric amount). In this embodiment, theglycerolysis reactors used as elements 134 and 136 are configured as twoheated, continuous stirred tank reactors in series. In these vessels,the mixture of glycerin and grease (containing free fatty acids) isagitated to keep the two immiscible fluids in intimate contact.

In a preferred embodiment, mixing is provided by an agitator. Underthese conditions, the free fatty acids are converted into glycerides(mono-, di-, or triglycerides) with the production of water. The wateris vented as vapor and removed from the system together with any waterthat was initially present in the feedstock in water vapor vent 35. Thefree fatty acid content of the reactor effluent stream in this preferredembodiment of the invention can consistently be maintained at less than0.5 percent w/w.

Because of the corrosive nature of free fatty acids, the glycerolysisreactor is preferably constructed of materials resistant to organicacids.

The effluent from glycerolysis reaction subsystem 126 contains mono-,di-, and triglycerides and residual fatty acids. The glycerolysisreaction effluent is introduced to alkali catalyzed transesterificationsubsystem 142 which preferably comprises methanol metering apparatus144, potassium hydroxide metering apparatus 146, methoxide additionapparatus 148 and first transesterification reactor 150 and secondtransesterification reactor 151 in which the glycerides undergotransesterification.

In transesterification reaction subsystem 142, the glycerides aretransesterified with an alkali catalyst and a simple alcohol having 1 to5 carbons. In a preferred embodiment, the alkali catalyst is potassiumhydroxide and the alcohol is methanol. The residual free fatty acids aresaponified consuming a molar quantity of alkali catalyst about equal tothe number of moles of free fatty acid present.

The transesterification reaction is preferably catalyzed by potassiummethoxide, which is formed from the addition of potassium hydroxide tomethanol. The amount of potassium hydroxide added is preferablyequivalent to 0.5 percent to 2.0 percent w/w of the glycerides presentin the feed solution. The methanol and catalyst are combined and addedto the solution of glycerides coming from the glycerolysis reactors bymethoxide addition apparatus 148.

A 200 percent stoichiometric excess of methanol based upon the number ofmoles of fatty acids available in the glycerides is added to thereaction mixture. Upon entering each transesterification reactor 150 and151, the two-phase system undergoes vigorous mixing.

Preferably, the reaction temperature is held between about 25° C. andabout 65° C. At this temperature, the miscibility of the phases islimited and mixing is required to achieve a high conversion rate. Theresidence time required is dependent on glyceride composition of thefeed (between mono-, di- and triglycerides), temperature, catalystconcentration and mass transfer rate.

Thus, agitation intensity is preferably considered in selecting aresidence time. Typically, the residence time required for greater than(>) 99 percent conversion of glycerides to alkyl esters is 20 to 30minutes.

In the transesterification reactor, the presence of potassium hydroxide,methanol, and fatty acid esters can be corrosive. In a preferredembodiment, at least two continuous stirred tank reactors in series areused. Suitable resistant materials are preferably chosen for thereactors.

The effluent from transesterification subsystem 142 may be introduced tophase separation subsystem 52 which comprise phase separation tank 54 inwhich a light phase (for instance, specific gravity 0.79−0.88) isseparated from a heavy phase (for instance, specific gravity 0.90−1.2).The effluent streams from the phase separator are a light phase fattyacid alkyl esters comprised of methanol and alkyl esters (biodiesel), afraction of the excess alcohol and some impurities, and a heavy phase(crude glycerin) containing glycerin, alcohol, FAAEs, soaps, alkalicatalyst, a trace of water and some impurities.

Phase separation unit 54 is preferably a conventional liquid/liquidseparator, capable of separating of the heavy phase from the lightphase. Suitable phase separation units include commercially availableequipment, including continuous clarifier 54.

In biodiesel purification subsystem 56, excess methanol and high-boilingimpurities may be separated from the fatty acid methyl esters in thelight phase in fractionation column 58 and methanol collected for reuse.Preferably, purifying the fatty acid methyl ester rich stream subsystem56 further comprises a fatty acid alkyl ester distillation column 58 forseparating the fatty acid alkyl ester rich stream into a bottomsfraction, an overhead fraction comprising primarily methanol, and a sidestream fraction comprising a fatty acid alkyl ester product.

Preferably, the bottoms fraction produced by distillation column 58comprises impurities, and unsaponifiable materials, monoglycerides,diglycerides, triglycerides and fatty acids. Preferably, the fatty acidmethyl ester product produced by distillation column 58 in FIG. 2 meetsASTM specification D 6751.

Preferably, the overhead fraction produced by distillation column 58comprises essentially methanol. Preferably, distillation column 58 isoperated under pressure below about 2 pounds per square inch absoluteand at a temperature in the range of about 180° C. to about 280° C. Morepreferably, distillation column 58 is operated under pressure in therange of about 0.1 pounds per square inch absolute to about 2 pounds persquare inch absolute and at a temperature in the range of about 180° C.to about 230° C. Preferably, distillation column 58 contains highefficiency structured packing material.

The heavy phase separated in phase separation tank 54 is preferablytreated in catalyst separation subsystem 62 comprising a mineral acid(such as phosphoric acid) addition apparatus 64, catalyst precipitationreactor 66, catalyst precipitation reactor effluent filter 70 in whichwashing with methanol 68 occurs before the potassium phosphateprecipitate 171 is removed from the filter, filtrate separation tank 72,pH neutralization tank and free fatty acid recycling apparatus 76.

In catalyst separation subsystem 62, the crude glycerin phase is pumpedto a catalyst precipitation reactor where a mineral acid 64 is added.Preferably, the amount of acid added is a molar quantity equal to themolar quantity of alkali catalyst used in the transesterificationreaction. The product of the reaction is an insoluble salt that can beseparated as a solid. In addition to forming an insoluble salt, the acidconverts soaps formed in transesterification reaction subsystem 142 tofree fatty acids.

In a preferred embodiment, potassium hydroxide is used as thetransesterification catalyst, and the precipitation reaction usesphosphoric acid to form monobasic potassium phosphate. This salt is notsoluble in this system and can be removed by simple filtration. As thepotassium phosphate salt is filtered in catalyst precipitation reactoreffluent filter 70, methanol 68 is used to wash glycerin and otherprocess chemicals off of the precipitate.

The filtrate from catalyst precipitation reactor effluent filter 70 issent to another phase separation operation where two liquid phases formand separate according to their relative specific gravities in filtrateseparation tank 72. Glycerin, water, impurities and most of the methanolreport to the bottom or heavy phase, while fatty acid alkyl ester, somealcohol and fatty acids report to the top, or light phase. The lightphase is combined with the light phase from the previous phaseseparation subsystem (subsystem 52) and sent to the fractionation column58. The heavy phase is sent to a reaction operation where any residualacid is neutralized in pH neutralization reactor 74 by adding a smallamount of caustic. In a preferred embodiment, this is performed in acontinuous stirred tank reactor.

Following pH neutralization reactor 74, the crude glycerin phase is sentto the glycerin refining subsystem 80, where the methanol and water areseparated and collected for further purification and the glycerin isseparated from the high boiling impurities. In a preferred embodiment,glycerin separation is performed in glycerin distillation orfractionation column 84 with a glycerin side draw. The distilledglycerin may further be treated in glycerin decolorization column 86 inwhich activated carbon is used to remove color and odor from thedistilled glycerin.

The methanol recovered from the distillation column contains traceamounts of water and is therefore considered a “wet” methanol streamthat must be purified prior to reuse in the process in methanolpurification subsystem 88. This “wet” methanol stream is collected andpurified by distillation in methanol purification column 90 before beingpumped back into the inventory storage tanks.

The distilled glycerin stream is then subjected to decolorization anddeodorization through activated carbon bed 86. The feed enters thecolumn from the bottom and is allowed to flow upwards through theactivated carbon bed resulting in a colorless, solventless and salt freeglycerin that is >95 percent pure.

Glycerin recycling pump 92 may be used to recycle glycerin to glycerinaddition apparatus 28. Methanol recycling apparatus 94 is preferablyused to recycle methanol to methanol metering apparatus 144.

Biodiesel is then delivered to its market in biodiesel delivery vehicle96 and glycerin is delivered to its market in glycerin delivery vehicle98.

The process may also consist of refinements to increase the yield ofproduction of biodiesel. FIG. 7 illustrates the option of increasing theyield in the production of biodiesel by further treatment of byproductstream 358, depending to a large extent on its relative concentration offatty acid alkyl esters, glycerides, and free fatty acids in by-productstream 358. As illustrated, a portion of by-product stream 358 may betreated in biodiesel refining step 8. As shown in FIG. 7, fatty acidalkyl ester enriched stream 351A of by-product fuel stream 351 isredirected to biodiesel refining stage 8 for further recovery of fattyacid alkyl esters. Stream 358, when containing significant portions ofglycerides, may further be introduced into transesterification reactor 4or esterification reactor 2. As illustrated, a fraction of by-productstream 358 is introduced as stream 351C into transesterification reactor4. Alternatively, stream 351D, when containing higher free fatty acidcontent is preferably introduced into esterification reactor 2.

In FIG. 8, a portion of by-product stream 358, represented as stream351, may first be separated, preferably in non-evaporative separator370, as fatty acid alkyl ester rich stream 371 and/or glyceridesenriched rich stream 376 and/or free fatty acids enriched stream 374.The fraction containing low free fatty acid content may then beintroduced as stream 376 into transesterification reactor 4 and stream374 containing higher free fatty acid content may be introduced intoesterification reactor 2. Suitable non-evaporative separation techniquesthat may be used are freeze crystallization, steam stripping orliquid-liquid separation.

Increased yield of biodiesel may further result by the use of a seconddistillation reactor or non-evaporative separator in biodiesel refiningstage 8. As shown in FIG. 9, a fatty acid rich stream, such as enrichedstream 323 separated from the transesterification effluent stream in1^(st) phase separation 320, is introduced to heat exchanger 405 andintroduced via pump 406 into flash drum 410. Typical operatingtemperature range for flash drum 410 is from about 60° C. to about 205°C., more typically about 140° C., and typical operating pressure is fromabout 1 pound per square inch absolute to about 15 pounds per squareinch absolute, more typically about 5 pounds per square inch absolute.Vapor 412 is removed and the liquid stream 411 is then pumped throughpump 415 into distillation column 420. In a preferred embodiment, asdiscussed above, distillation column 420 is a reactive distillationcolumn. Overhead fraction 422 enters heat exchanger 440 and exits thesystem in vapor form, principally as excess alcohol, as stream 442.Condensate 441A exiting heat exchanger 440 exits the system and liquidstream 441B re-enters the distillation column. The bottoms fraction 421from distillation column 420 is principally the fatty acid alkyl esterrich stream and may then be introduced into reboiler 430 where it iseither further separated as vapor stream 432 in distillation column 420or exits as biodiesel stream 431A. Biodiesel stream 431A consistsprincipally of fatty acid alkyl esters, glycerides and a trace amount ofglycerin and, depending on the acidity upstream, some fatty acids. Thisstream may further be subjected to a second distillation in distillationcolumn 450, via holding tank 440, to render purified biodiesel stream350C and by-product (fuel) stream 350A. In a preferred embodiment,distillation column 450 is either one or more wiped film evaporators orfalling film evaporators known in the art. The temperature in the seconddistillation column 450 is approximately the same as the temperature indistillation column 420. In an alternative embodiment, shown in FIG. 10,a portion of by-product (fuel) stream 350A may be re-introduced tosecond distillation column 450 via holding tank 440.

The second distillation procedure may occur in one or more distillationcolumns. For instance, a single wiped film evaporator or falling filmevaporator may be used. Further, multiple wiped film evaporators orfalling film evaporators in parallel or series may be used. Residencetime of the biodiesel stream in the wiped film evaporator and failingfilm evaporator is generally short.

The wiped film evaporator consists of internal rotating distributorplates which serve to evenly disperse the biodiesel at the top of theheated plate of the evaporator to the interior surfaces of a heatedcylindrical shell. Wiper blades then spread, agitate and move thebiodiesel downwards along the heated shell in rapid time while fattyacid alkyl esters are quickly evaporated and re-condensed on a cooledsurface, typically at the center of the evaporator. With this particularconfiguration, the purified biodiesel stream then exits the bottom ofthe center of the evaporator, and the byproduct (fuel) stream exits thebottom of the outer perimeter of the evaporator.

The falling film evaporator consists of an outer shell filled with steamor other heating media and vertical, parallel tubes through which thebiodiesel falls. The flow of biodiesel is controlled such that thebiodiesel creates a film along the inner tube walls, which progressesdownwards while the biodiesel is selectively evaporated from the liquid.Separation between biodiesel vapors and the residual liquid typicallyconsisting of a mixture of glycerides, fatty acids and some unevaporatedfatty acid alkyl esters occurs in the tubes. The biodiesel vapor isliquefied in a cooled condenser and recovered.

As in distillation column 420, these second distillation columns 450,are typically operated at a pressure below about 250 torr absolute andat a temperature in the range of about 150° C. to about 320° C. Morepreferably, distillation column 450 is operated at a pressure in therange of about 0.1 torr absolute to about 2 torr absolute and at atemperature in the range of about 180° C. to about 230° C.

FIG. 11 presents a further embodiment of the invention wherein theby-product (fuel) stream 350A is introduced to separator 370. Separator370 is preferably a non-evaporative separator. A fatty acid alkyl esterenriched stream 371 may be separated from a stream 372 enriched inglycerides and/or free fatty acids in separator 370. The fatty acidalkyl ester enriched stream 371 may then be re-introduced to seconddistillation column 450 via holding tank 440 for further separation intopurified biodiesel. Stream 372 enriched in glycerides and/or free fattyacids may then be re-introduced into transesterification reactor 4 andesterification reactor 2.

FIG. 12 presents another embodiment wherein the fatty acid alkyl esterrich stream 371 may be branched and introduced as stream 371A with thepurified biodiesel stream 350C. In addition, a portion of stream 371 maybe re-introduced to the second distillation column 450. Further, FIG. 12illustrates the option of introducing either a portion or all ofbiodiesel stream 431A, as 452, from first distillation column 420 intoseparator 370 for separation into a fatty acid alkyl ester rich streamand a glyceride and/or free fatty acid rich stream 372. The glycerideand/or free fatty acid enriched stream 372 may then be re-introducedinto the transesterification reactor 4 and/or esterification reactor 2.

With respect to the above description then, it is to be realized thatthe optimum dimensional relationships for the parts of the invention, toinclude variations in size, materials, shape, form, function and mannerof operation, assembly and use, are deemed readily apparent and obviousto one skilled in the art, and all equivalent relationships to thoseillustrated in the drawings and described in the specification areintended to be encompassed by the present invention.

Therefore, the foregoing is considered as illustrative only of theprinciples of the invention. Further, since numerous modifications andchanges will readily occur to those skilled in the art, it is notdesired to limit the invention to the exact construction and operationshown and described, and accordingly, all suitable modifications andequivalents may be resorted to, falling within the scope of theinvention.

EXAMPLES Example No. 1

Rendered yellow grease with a free fatty acid concentration of 20percent by weight and 2 percent moisture, impurities and unsaponifiables(MIU) was fed to continuous stirred tank glycerolysis reactors at 100pounds per minute (lbs/min). The grease was filtered and titratedintermittently as it was fed to the glycerolysis reactor. Glycerin wasadded at a rate of 13 lbs/min. The temperature of the grease andglycerin mixture was raised to 210° C. as it was fed into the first ofthe glycerolysis continuous stirred tank reactors. In the reactor, thepressure was reduced to 2 psia and the temperature was maintained at210° C. The vessel was fitted with a high intensity agitator to keep theimmiscible liquids in contact. Water vapor produced by the reaction wasremoved through vents in the reactor headspace. The residence time ineach of the glycerolysis reactors was 2.5 hours. The conversion of fattyacids to glycerides in the first vessel was 85 percent. The fatty acidconcentration leaving the second reactor was maintained at 0.5 percentw/w.

The product from the glycerolysis reactors was cooled to 50° C. and fedcontinuously to the transesterification reactors in which a solution ofpotassium hydroxide in methanol was added. The potassium hydroxide wasadded at a rate of 1.1 lbs/min and mixed with 22 lbs/min of methanol.The transesterification took place in two continuous stirred tankreactors in series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters, a small amount ofunreacted glycerides and a small concentration of the unreacted methanolfloated to the top. The glycerin, the majority of the unreactedmethanol, some fatty acid methyl esters, potassium hydroxide and soapssank to the bottom.

The bottom, or heavy phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterification stepwas reacted with 1.96 lbs/min phosphoric acid. The soaps converted tofree fatty acids and the potassium hydroxide was neutralized. Theproduct of this acidification was monobasic potassium phosphate, whichwas not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out and thefiltrate was fed to a second phase separation tank where the fatty acidmethyl esters and free fatty acids present in the filtrate floated tothe top and the glycerin and methanol sank to the bottom. The top, orlight, phase was mixed with the light phase from the first phaseseparation tank and fed to the fatty acid methyl ester fractionationcolumn. The pH of the heavy phase was adjusted back to 7.5 withpotassium hydroxide and fed to the glycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and18 lbs/min of glycerin. The glycerin produced was more than 95 percentpure with non-detectable concentrations of salts and methanol. Thisglycerin stream was split into two streams: 13 lbs/min was recycled backto the glycerin feed tank for the glycerolysis reaction and 5 lbs/minwas pumped through the decolorization column and collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2 lbs/min of methanol was recovered and 92lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) were produced.

Example No. 2

Fancy bleachable inedible tallow with a free fatty acid concentration of4 percent by weight and 0.5 percent MIU (moisture, impurities andunsaponifiables) was fed to a continuous stirred tank reactor at 100lbs/min. The grease was filtered and titrated continuously as it was fedto the glycerolysis reactors. Glycerin was added at a rate of 2.6lbs/min. The temperature of the grease and glycerin mixture was raisedto 210° C. as it was fed into the first of the glycerolysis continuousstirred tank reactors. In the reactor the pressure was reduced to 2 psiaand the temperature was maintained. The vessel was fitted with anagitator to keep the immiscible liquids in contact. Water vapor producedby the reaction was removed through vents in the reactor headspace. Theresidence time in each of the glycerolysis reactors was 2.5 hours. Theconversion of fatty acids to glycerides in the first vessel was 92percent. The fatty acid concentration leaving the second reactor wasmaintained at 0.5 percent by weight.

The product from the glycerolysis reactors was cooled to 50° C. and fedto the transesterification reactors in which a solution of potassiumhydroxide in methanol was added. The potassium hydroxide was added at arate of 1.0 lbs/min and mixed with 22 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and a smallconcentration of the unreacted methanol floated to the top. Theglycerin, the majority of the unreacted methanol, some fatty acid methylesters, potassium hydroxide and soaps sank to the bottom.

The bottom, or heavy phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterificationoperation was reacted with 1.79 lbs/min phosphoric acid. The soapsconverted back to free fatty acids and the potassium hydroxide wasneutralized. The product of this acidification was monobasic potassiumphosphate, which was not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out and thefiltrate was fed to a second phase separation tank where the fatty acidmethyl esters and free fatty acids floated to the top and the glycerinand methanol sank to the bottom. The top, or light, phase was mixed withthe light phase from the first phase separation tank and fed to thefatty acid methyl ester fractionation column. The pH of the heavy phasewas adjusted to 7.8 with 0.1 lbs/min potassium hydroxide and fed to theglycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and10.2 lbs/min of glycerin. The glycerin produced was more than 95 percentpure with non-detectable concentrations of salts and methanol. Theglycerin stream was split into two streams: 2.6 lbs/min was recycledback to the glycerin feed tank for the glycerolysis reaction and 7.6lbs/min was collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column in which 2.1 lbs/min of methanol was recovered and93 lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) was produced.

Example No. 3

Degummed, food-grade soybean oil with a free fatty acid concentration of0.5 percent by weight and 0.5 percent MIU (moisture, impurities andunsaponifiables) was fed to a conditioning chamber at 100 lbs/min. Thegrease was filtered and titrated continuously as it was transferred fromthe feedstock conditioner. Due to the low concentration of free fattyacids, the glycerolysis section of the process was bypassed when usingthis feedstock.

The fatty acid concentration entering the transesterification reactorswas 0.5 percent by weight. The potassium hydroxide was added at a rateof 1.0 lbs/min and mixed with 22 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and a smallconcentration of the unreacted methanol floated to the top. Theglycerin, the majority of the unreacted methanol, some fatty acid methylesters, potassium hydroxide and soaps sank to the bottom.

The bottom, or heavy, phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterificationoperation was reacted with 1.76 lbs/min phosphoric acid. The pH of thesolution was decreased, and the product of this acidification wasmonobasic potassium phosphate, which was not soluble in this system.

The precipitate was filtered out at 2.2 lbs/min and the filtrate was fedto a phase separation tank in which the fatty acid methyl esters andfree fatty acids floated to the top and the glycerin and methanol sankto the bottom. The top, or light, phase was mixed with the light phasefrom the first phase separation tank and fed to the fatty acid methylester fractionation column. The heavy phase was transferred to anothertank and the pH was adjusted to 7.4 with 0.1 lbs/min potassiumhydroxide. Then, the glycerin/methanol mixture was fed to the glycerinfractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and8.5 lbs/min of glycerin. The glycerin produced had a purity greater than95 percent with non-detectable concentrations of salts and methanol. Theglycerin was collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2.1 lbs/min of methanol was recovered and 93lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) were produced.

Example No. 4

Rendered trap grease with a free fatty acid concentration of 68 percentby weight and 5% MIU (moisture, impurities and unsaponifiables) was fedto the invention at 100 lbs/min. The grease was filtered and titratedcontinuously as it was fed to the glycerolysis reactors. Glycerin wasadded at a rate of 44 lbs/min. The temperature of the grease andglycerin mixture was raised to 210° C. as it was fed into the first ofthe glycerolysis continuous stirred tank reactors. In the reactor, thepressure was reduced to 2 psia and the temperature was maintained. Watervapor produced by the reaction was removed through vents in the reactorheadspace. The residence time in each of the glycerolysis reactors was3.5 hours. The conversion of fatty acids to glycerides in the firstvessel was 87 percent. The fatty acid concentration leaving the secondreactor was maintained at 0.5 percent by weight.

The product from the glycerolysis reactors was cooled to 50° C. and fedto the transesterification reactors where a solution of potassiumhydroxide in methanol was added. The potassium hydroxide was added at arate of 1.4 lbs/min and mixed with 21 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and 10 percent of theunreacted methanol floated to the top and the glycerin, the majority ofthe unreacted methanol, some fatty acid methyl esters, potassiumhydroxide and soaps sank to the bottom.

The bottom, or heavy, phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterificationoperation was reacted with 2.45 lbs/min phosphoric acid. The soapsconverted back to free fatty acids and the potassium hydroxide wasneutralized. The product of this acidification was monobasic potassiumphosphate, which was not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out at 3.1lbs/min and the filtrate was fed to a second phase separation tank wherethe fatty acid methyl esters and free fatty acids floated to the top andthe glycerin and methanol sank to the bottom. The top, or light, phasewas mixed with the light phase from the first phase separation tank andfed to the fatty acid methyl esters fractionation column. The pH of theheavy phase was adjusted back to 7.3 with 0.14 lbs/min potassiumhydroxide and fed to the glycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and40 lbs/min of glycerin. The glycerin produced had a purity greater than95 percent with non-detectable concentrations of salts and methanol.This glycerin stream was recycled back to the glycerin feed tank for theglycerolysis reaction and an additional 4 lbs/min of fresh glycerin wasadded to the glycerin feed tank to provide enough glycerin feed for theglycerolysis reaction.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2.1 lbs/min of methanol was recovered and 91lbs/min of fatty acid methyl esters meeting ASTM D 6751-02 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels) were produced.

Example No. 5

Rendered brown grease with a free fatty acid concentration of 37 percentby weight and 5 percent MIU (moisture, impurities and unsaponifiables)was fed to the invention at 100 lbs/min. The grease was filtered andtitrated continuously as it was fed to the glycerolysis reactors.Glycerin was added at a rate of 24 lbs/min. The temperature of thegrease and glycerin mixture was raised to 210° C. as it was fed into thefirst of the glycerolysis continuous stirred tank reactors. In thereactor, the pressure was reduced to 2 psia and the temperature wasmaintained. The vessel was fitted with an agitator to keep theimmiscible liquids in contact. Water vapor produced by the reaction wasremoved through vents in the reactor headspace. The residence time ineach of the glycerolysis reactors was 3.0 hours. The conversion of fattyacids to glycerides in the first vessel was 90 percent. The fatty acidconcentration leaving the second reactor was maintained at 0.5 percentby weight.

The product from the glycerolysis reactors was cooled to 50° C. and fedto the transesterification reactors where a solution of potassiumhydroxide in methanol was added. The potassium hydroxide was added at arate of 1.2 lbs/min and mixed with 21 lbs/min of methanol. Thetransesterification took place in two continuous stirred tank reactorsin series, each with a two-hour residence time.

The transesterified product was then fed to a phase separation tankwhere the majority of the fatty acid methyl esters and 10 percent of theunreacted methanol floated to the top. The glycerin, the majority of theunreacted methanol, some fatty acid methyl esters, potassium hydroxideand soaps sank to the bottom.

The bottom, or heavy, phase was sent to an acidification reactor wherethe potassium hydroxide catalyst added in the transesterification wasreacted with 2.13 lbs/min phosphoric acid. The soaps converted back tofree fatty acids and the potassium hydroxide was neutralized. Theproduct of this acidification was monobasic potassium phosphate, whichis not soluble in this system.

The monobasic potassium phosphate precipitate was filtered out at 2.7lbs/min and the filtrate was fed to a second phase separation tank wherethe fatty acid methyl esters and free fatty acids floated to the top andthe glycerin and methanol sank to the bottom. The top, or light, phasewas mixed with the light phase from the first phase separation tank andfed to the fatty acid methyl ester fractionation column. The pH of theheavy phase was adjusted to 7.5 with 0.12 lbs/min potassium hydroxideand fed to the glycerin fractionation column.

The glycerin fractionation column recovered 10 lbs/min of methanol and25.2 lbs/min of glycerin. The glycerin produced had a purity greaterthan 95 percent with non-detectable concentrations of salts andmethanol. This glycerin stream was split into two streams: 24 lbs/minwas recycled back to the glycerin feed tank for the glycerolysisreaction, and 1.2 lbs/min was collected for market.

The two light phase streams were fed to the fatty acid methyl esterfractionation column where 2.0 lbs/min of methanol was recovered, and89.8 lbs/min of fatty acid methyl esters meeting ASTM D 6751-02(Standard Specification for Biodiesel Fuel (B100) Blend Stock forDistillate Fuels) were produced.

Example No. 6

A feedstock containing about 0.3 weight percent of free fatty acids andabout 99.3 weight percent of glycerides (the remainder being water andinsoluble and unsaponifiable solids), at a flow rate of about 40.9pounds per hour, was heated to 50° C. and added to a solution ofpotassium hydroxide (1 percent of the feedstock flow on a weight basis)in methanol (stoichiometric ratio of 2:1 methanol:bound fatty acids inglycerides). The transesterification took place in a single continuousstirred tank reactor with a ten-hour residence time.

The transesterification effluent stream flow rate was approximately 50.3pounds per hour and consisted of approximately 79 weight percent offatty acid methyl esters, 8 weight percent glycerin, 9 weight percentmethanol, 1.6 weight percent glycerides, with the remainder being water,insoluble and unsaponifiable solids, and soaps.

This stream was separated in a flow-through separator into a light phasestream and a heavy phase stream, the light phase stream having a flow of41.5 pounds per hour and a composition of approximately 94.26 weightpercent fatty acid methyl esters, 5.6 weight percent methanol, 0.09weight percent glycerides and 0.05 weight percent free glycerin.

Free glycerin concentrations in this and the other samples in thisexample were determined using an enzyme assay solution provided bySigma-Aldrich, Inc. of St. Louis, Mo. in a kit with product code BQP-02.With this kit, free glycerin was measured by coupled, enzymaticreactions that ultimately produce a quinoneimine dye that shows anabsorbance maximum at 540 nm. The absorbance peak was measured using aBausch & Lomb Spectronic 20 spectrophotometer.

The light phase stream was analyzed for glycerin and found to containapproximately 490 ppm glycerin by weight. The light phase stream wasintroduced into a reactive distillation column maintained at 260° C. ata pressure of 150 mmHg. The overhead vapor stream from the column wascondensed, producing a liquid stream with a flow rate of about 2.1pounds per hour, consisting primarily of methanol with a glycerincontent of 135 ppm. The bottoms liquid stream, having a flow rate ofapproximately 39.3 pounds per hour, consisted of approximately 98.5weight percent fatty acid methyl esters, 1.5 weight percent glycerides,and only 3 ppm glycerin. The reactive distillation referenced in thisparagraph is schematically displayed as FIG. 6.

The gravimetric flow rates calculated using these analyses of freeglycerin in the feed to the column versus in the overhead and bottomsstreams indicated that about 98 percent of the glycerin was reacted intoother moieties in the distillation column rather than simply flowingeither to the overhead or bottoms streams.

This bottoms liquid stream was further refined to produce a biodieselstream of fatty acid methyl esters meeting ASTM D 6751-06 S15 (StandardSpecification for Biodiesel Fuel (B100) Blend Stock for DistillateFuels).

From the foregoing, it will be observed that numerous variations andmodifications may be effected without departing from the true spirit andscope of the novel concepts of the invention.

1. A process for the production of purified biodiesel from glyceridescomprising: (A) reacting glycerides with at least one alcohol in atransesterification reactor to produce a fatty acid alkyl ester stream;(B) separating a first biodiesel stream from the fatty acid alkyl esterstream by a first distillation or non-evaporative separation; and (C)subjecting at least a portion of the first biodiesel stream to a seconddistillation or non-evaporative separation to render a purified secondbiodiesel stream and a by-product fuel stream.
 2. The process of claim1, wherein the first biodiesel stream is separated from the fatty acidalkyl ester stream in at least one wiped film evaporator.
 3. The processof claim 2, wherein at least a portion of the by-product fuel stream isre-introduced into the at least one wiped film evaporator.
 4. Theprocess of claim 1, wherein the first biodiesel stream is separated fromthe fatty acid alkyl ester stream in at least one falling filmevaporator.
 5. The process of claim 4, wherein at least a portion of theby-product fuel stream is re-introduced into the at least one fallingfilm evaporator.
 6. The process of claim 1, further comprisingseparating from at least a portion of the first biodiesel stream a freefatty acid and/or glyceride enriched stream.
 7. The process of claim 6,wherein the free fatty acids of the free fatty acid and/or glycerideenriched stream are converted to glycerides.
 8. The process of claim 7,wherein the glycerides are introduced to the transesterification reactorof step (A).
 9. The process of claim 6, wherein the free fatty acidand/or glyceride enriched stream is separated from the first biodieselstream by freeze crystallization.
 10. The process of claim 6, wherein afatty acid ester enriched stream is further separated from at least aportion of the first biodiesel stream.
 11. The process of claim 10,wherein the fatty acid ester enriched stream is subjected to the seconddistillation or non-evaporative separation step of step (C).
 12. Theprocess of claim 2, wherein the at least one wiped film evaporator is atleast two parallel wiped film evaporators.
 13. The process of claim 4,wherein the at least one falling film evaporator is at least twoparallel falling film evaporators.
 14. The process of claim 1, whereinthe by-product fuel stream is further separated into a stream enrichedin free fatty acids and/or glycerides.
 15. The process of claim 14,wherein the free fatty acid and/or glyceride enriched stream isseparated from the by-product fuel stream by freeze crystallization. 16.The process of claim 1, wherein the at least one alcohol is a C₁-C₅alcohol.
 17. The process of claim 16, wherein the C₁-C₅ alcohol ismethanol.
 18. The process of claim 1, wherein the process is continuous.19. In a process for the production of biodiesel from glycerides whereinglycerides are reacted with an alcohol to produce a fatty acid alkylester stream which is separated into a biodiesel stream and a by-productfuel stream, the improvement comprising separating a second biodieselstream and a second by-product fuel stream from the by-product fuelstream.
 20. The process of claim 19, wherein the second biodiesel streamand second by-product fuel stream are separated from the by-product fuelstream by freeze crystallization.
 21. A process for the production ofpurified biodiesel from glycerides comprising: (A) reacting glycerideswith at least one alcohol to produce a fatty acid alkyl ester stream;(B) separating by distillation or non-evaporative separation a biodieselstream and a by-product fuel stream from the fatty acid alkyl esterstream; and (C) separating from the by-product fuel stream a secondbiodiesel stream and a fatty acid and/or glyceride enriched stream. 22.The process of claim 21, wherein the second biodiesel stream and fattyacid and/or glyceride enriched stream are separated from the by-productfuel stream by freeze crystallization.
 23. A process for the productionof purified biodiesel from glycerides comprising: (A) reacting theglycerides with at least one alcohol in a transesterification reactor toproduce a fatty acid alkyl ester stream; (B) purifying the fatty acidalkyl ester stream by distillation or non-evaporative separation andseparating therefrom a first biodiesel stream; (C) separating from atleast a portion of the first biodiesel stream a stream enriched in freefatty acids and/or glycerides; (D) introducing the glycerides of step(C) into the transesterification reactor; and (E) recovering biodieseltherefrom.
 24. The process of claim 23, wherein the free fatty acids inthe stream enriched in free fatty acids and/or glycerides of step (C)are further converted to glycerides.
 25. The process of claim 23,wherein the free fatty acids in the glyceride and/or free fatty acidenriched stream of step (D) are further converted to glycerides.
 26. Theprocess of claim 23, wherein the second distillation is conducted in atleast one wiped film evaporator.
 27. The process of claim 23, whereinthe second distillation is conducted in at least one falling filmevaporator.
 28. The process of claim 23, wherein the at least onealcohol is a C₁-C₅ alcohol.
 29. The process of claim 28, wherein theC₁-C₅ alcohol is methanol.
 30. The process of claim 23, wherein theprocess is continuous.
 31. A process for the production of purifiedbiodiesel from glycerides comprising: (A) reacting glycerides with atleast one alcohol in a transesterification reactor to produce a fattyacid alkyl ester stream; (B) separating from the fatty acid alkyl esterstream a biodiesel stream and a by- product fuel stream; (C) introducingthe glycerides from the by-product fuel stream into thetransesterification reactor; and (D) recovering biodiesel therefrom. 32.The process of claim 31, wherein the at least one alcohol is a C₁-C₅alcohol.
 33. The process of claim 32, wherein the C₁-C₅ alcohol ismethanol.
 34. The process of claim 31, wherein the process iscontinuous.
 35. In a process for the production of biodiesel wherein afatty acid alkyl ester stream is separated into a first biodiesel streamand a first by-product fuel stream, the improvement comprisingseparating from at least a portion of the first by-product fuel stream astream enriched in free fatty acids and/or glycerides from which asecond biodiesel stream is produced.
 36. A process for producing apurified biodiesel comprising: (A) separating a first biodiesel streamfrom a fatty acid alkyl ester enriched stream; (B) introducing at leasta portion of the first biodiesel stream to at least one wiped filmevaporator or at least one falling film evaporator and separating asecond biodiesel stream and a by-product fuel stream therefrom; (C)separating from the by-product fuel stream a fatty acid alkyl esterenriched stream and a free fatty acid/glyceride enriched stream; (D)recovering purified biodiesel from the fatty acid alkyl ester enrichedstream; (E) producing a fatty acid alkyl ester enriched stream from thefree fatty acid/glyceride enriched stream from step (C); and (F)recovering purified biodiesel from the stream of step (E).
 37. Theprocess of claim 36, wherein at least a portion of the fatty acid alkylester enriched stream of step (C) is introduced into at least one wipedfilm evaporator.
 38. The process of claim 37, wherein at least one wipedfilm evaporator is at least two parallel wiped film evaporators.
 39. Theprocess of claim 36, wherein at least a portion of the fatty acid alkylester enriched stream of step (C) is introduced into at least onefalling film evaporator.
 40. The process of claim 39, wherein at leastone falling film evaporator is at least two parallel falling filmevaporators.